U.S. patent number 10,865,167 [Application Number 16/331,828] was granted by the patent office on 2020-12-15 for hydrocracking process.
This patent grant is currently assigned to SABIC GLOBAL TECHNOLOGIES B.V.. The grantee listed for this patent is SABIC GLOBAL TECHNOLOGIES B.V.. Invention is credited to Luis Aramburo, Ashim Kumar Ghosh, Alla Khanmamedova, Cong Nguyen, Scott A. Stevenson, David L. Sullivan.
United States Patent |
10,865,167 |
Ghosh , et al. |
December 15, 2020 |
Hydrocracking process
Abstract
A process for hydrocracking 2,4-dimethylpentane and/or
2,2,3-trimethylbutane can comprise: contacting a hydrocracking feed
stream in the presence of hydrogen with a hydrocracking catalyst,
wherein the hydrocracking feed stream comprises at least 0.5 wt %
of 2,4-dimethylpentane and/or 2,2,3-trimethylbutane, based upon a
total weight of the hydrocracking feed stream; and wherein the
hydrocracking catalyst comprises a medium pore zeolite having a
pore size of 5-6 A and a silica to alumina molar ratio of 20-75;
preferably the hydrocracking catalyst comprises a medium pore
zeolite having a pore size of 5-6 A and a silica to alumina molar
ratio of 20-75 and a large pore zeolite having a pore size of 6-8 A
and a silica to alumina molar ratio of 10-80, wherein the
hydrogenation metal is deposited on the medium pore zeolite and the
large pore zeolite.
Inventors: |
Ghosh; Ashim Kumar (Sugar Land,
TX), Khanmamedova; Alla (Sugar Land, TX), Stevenson;
Scott A. (Sugar Land, TX), Aramburo; Luis (Geleen,
NL), Sullivan; David L. (Little Ferry, NJ),
Nguyen; Cong (Sugar Land, TX) |
Applicant: |
Name |
City |
State |
Country |
Type |
SABIC GLOBAL TECHNOLOGIES B.V. |
Bergen op Zoom |
N/A |
NL |
|
|
Assignee: |
SABIC GLOBAL TECHNOLOGIES B.V.
(Bergen op Zoom, NL)
|
Family
ID: |
1000005243174 |
Appl.
No.: |
16/331,828 |
Filed: |
September 7, 2017 |
PCT
Filed: |
September 07, 2017 |
PCT No.: |
PCT/IB2017/055397 |
371(c)(1),(2),(4) Date: |
March 08, 2019 |
PCT
Pub. No.: |
WO2018/047093 |
PCT
Pub. Date: |
March 15, 2018 |
Prior Publication Data
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|
|
Document
Identifier |
Publication Date |
|
US 20190375696 A1 |
Dec 12, 2019 |
|
Foreign Application Priority Data
|
|
|
|
|
Sep 12, 2016 [EP] |
|
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16188306 |
|
Current U.S.
Class: |
1/1 |
Current CPC
Class: |
B01J
35/1057 (20130101); B01J 29/44 (20130101); C07C
4/06 (20130101); B01J 23/42 (20130101); B01J
35/1061 (20130101); B01J 29/22 (20130101); B01J
35/0066 (20130101); C07C 2529/44 (20130101); C07C
2529/40 (20130101); C07C 2523/42 (20130101); C07C
2523/44 (20130101); C07C 2523/10 (20130101); C07C
2529/18 (20130101); B01J 2229/186 (20130101); C07C
2521/12 (20130101) |
Current International
Class: |
C07C
4/06 (20060101); B01J 29/44 (20060101); B01J
35/00 (20060101); B01J 35/10 (20060101); B01J
23/42 (20060101); B01J 29/22 (20060101) |
References Cited
[Referenced By]
U.S. Patent Documents
Foreign Patent Documents
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105413741 |
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Mar 2016 |
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CN |
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9400409 |
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Jan 1994 |
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WO |
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0244306 |
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Jun 2002 |
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WO |
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2007055488 |
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May 2007 |
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WO |
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2010102712 |
|
Sep 2010 |
|
WO |
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2013182534 |
|
Dec 2013 |
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WO |
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2015189058 |
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Dec 2015 |
|
WO |
|
Other References
"Molecular Sieves"; Kirk-Othmer Encyclopedia of Chemical
Technology, vol. 16, pp. 811-853. cited by applicant .
International Search Report; International Application No.
PCT/IB2017/055397; International Filing Date: Sep. 7, 2017; dated
Jan. 2, 2018; 6 pages. cited by applicant .
Written Opinion; International Application No. PCT/IB2017/055397;
International Filing Date: Sep. 7, 2017; dated Jan. 2, 2018; 7
pages. cited by applicant .
China Office Action and Search Report for China Application No.
201780055434.3; Application Filing Date Mar. 8, 2019; dated Jun.
30, 2020; with English Translation, 20 pages. cited by
applicant.
|
Primary Examiner: Fadhel; Ali Z
Attorney, Agent or Firm: Cantor Colburn LLP
Claims
The invention claimed is:
1. A process of hydrocracking at least one of 2,4-dimethylpentane
and 2,2,3-trimethylbutane, comprising: contacting a hydrocracking
feed stream comprising C.sub.5-C.sub.12 hydrocarbons in the
presence of hydrogen with a hydrocracking catalyst under process
conditions including a temperature of 425-580.degree. C., a
pressure of 300-5000 kPa gauge and a Weight Hourly Space Velocity
of 0.1-30 to produce a hydrocracking product stream comprising LPG;
wherein the hydrocracking feed stream comprises at least 0.5 wt %
of 2,4-dimethylpentane and/or 2,2,3-trimethylbutane, based upon a
total weight of the hydrocracking feed stream; and wherein the
hydrocracking catalyst comprises a hydrogenation metal in an amount
of 0.010-0.30 wt % with respect to the total catalyst; and wherein
the hydrocracking catalyst comprises a medium pore zeolite having a
pore size of 5-6 .ANG. and a silica to alumina molar ratio of
20-75.
2. The process according to claim 1, wherein the total amount of
2,4-dimethylpentane and 2,2,3-trimethylbutane in the hydrocracking
feed stream is at least 1.0 wt %.
3. The process according to claim 1, wherein the hydrocracking
catalyst comprises 0.08 to 0.25 wt % hydrogenation metal, 15 wt %
to 25 wt % alumina, and a balance being the medium pore
zeolite.
4. The process according to claim 1, wherein the hydrocracking
catalyst is in the form of powder and is free from a binder.
5. The process according to claim 1, wherein the silica to alumina
molar ratio of the medium pore zeolite is in the range of
20-50.
6. The process according to claim 5, wherein the silica to alumina
molar ratio of the medium pore zeolite is in the range of 20 to
30.
7. The process according to claim 6, wherein the silica to alumina
molar ratio of the medium pore zeolite is in the range of 21 to
29.
8. The process according to claim 1, wherein a conversion of any
2,4-dimethylpentane is greater than or equal to 90%; and the
conversion of any 2,2,3-trimethylbutane is greater than or equal to
90%.
9. The process according to claim 1, wherein the hydrocracking
catalyst comprises at least 0.030 wt %, of the hydrogenating metal
in relation to the total weight of the catalyst.
10. The process according to claim 1, wherein the hydrocracking
catalyst comprises La and/or Ga.
11. The process according to claim 1, wherein the process comprises
separating BTX or benzene from the hydrocracking product
stream.
12. The process according to claim 1, wherein greater than or equal
to 95% of the C.sub.5-C.sub.12 hydrocarbons are cracked.
13. The process according to claim 1, wherein the hydrogenating
metal is at least one element selected from palladium and
platinum.
14. The process according to claim 1, wherein the hydrocracking
feed stream comprises benzene.
15. The process according to claim 1, wherein the hydrocracking
product stream comprises a greater amount of benzene compared to
the hydrocracking feed stream.
16. A process of hydrocracking at least one of 2,4-dimethylpentane
and 2,2,3-trimethylbutane, comprising: contacting a hydrocracking
feed stream comprising C.sub.5-C.sub.12 hydrocarbons in the
presence of hydrogen with a hydrocracking catalyst under process
conditions including a temperature of 425-580.degree. C., a
pressure of 300-5000 kPa gauge and a Weight Hourly Space Velocity
of 0.1-30 to produce a hydrocracking product stream; wherein the
hydrocracking feed stream comprises at least 0.5 wt % of
2,4-dimethylpentane and/or 2,2,3-trimethylbutane, based upon a
total weight of the hydrocracking feed stream; wherein the
hydrocracking catalyst comprises a hydrogenation metal in an amount
of 0.01-0.30 wt % with respect to the total catalyst; and wherein
the hydrocracking catalyst comprises a medium pore zeolite having a
pore size of 5-6 .ANG. and a silica to alumina molar ratio of 20-75
and a large pore zeolite having a pore size of 6-8 .ANG. and a
silica to alumina molar ratio of 10-80, wherein the hydrogenation
metal is deposited on the medium pore zeolite and the large pore
zeolite.
17. The process according to claim 16, wherein the zeolite in the
hydrocracking catalyst comprises 75-95 wt % of the medium pore
zeolite and 5-25 wt % of the large pore zeolite with respect to the
total amount of the zeolite.
18. The process according to claim 16, wherein the hydrocracking
catalyst has a deactivation rate of less than |-3.5.times.10.sup.-4
per hour|.
19. The process according to claim 16, wherein the medium pore
zeolite comprises a ZSM-5.
20. The process according to claim 16, wherein the large pore
zeolite comprises a mordenite.
Description
CROSS REFERENCE TO RELATED APPLICATIONS
This application is a 371 of International Application No.
PCT/IB2017/055397, filed Sep. 7, 2017, which is incorporated herein
by reference in its entirety, and which claims priority to European
Application No. 16188306.1, filed Sep. 12, 2016.
The present invention relates to a hydrocracking process.
It has been previously described in WO 02/44306 A1 and WO
2007/055488 A1 that aromatic hydrocarbon compounds and liquefied
petroleum gas (LPG) can be produced from a mixed hydrocarbon
feedstock having boiling points of 30-250.degree. C. Therefore a
hydrocarbon feedstock having boiling points of 30-250.degree. C.
and hydrogen is introduced to a reaction zone wherein said
hydrocarbon feedstock is converted in the presence of a catalyst to
aromatic hydrocarbon compounds abundant in BTX (benzene, toluene,
xylene) through hydrodealkylation and/or transalkylation and to
non-aromatic hydrocarbon compounds which are abundant in LPG
through hydrocracking and recovering the aromatic hydrocarbon
compounds and LPG, respectively, through gas-liquid separation and
distillation. The methods of WO 02/44306 A1 and WO 2007/055488
produce a product stream comprising a relatively high amount of
non-aromatic hydrocarbons that co-boil with BTX rendering it
impossible to produce chemical grade BTX without using solvent
extraction methods and a relatively high amount of fuel gas at the
expense of the LPG produced.
US2009/0272672 discloses a process for the catalytic
hydrodealkylation of C.sub.8-C.sub.13 alkyl aromatic compounds
mixed with C.sub.4-C.sub.10 aliphatic and cycloaliphatic products
which undergo aromatization and subsequent hydrodealkylation. In
this process, the hydrocarbons are treated with a ZSM-5 zeolite
having the silica to alumina (SiO.sub.2/Al.sub.2O.sub.3) molar
ratio of 5-100 modified by means of a platinum-molybdenum couple at
a temperature of 400 to 650.degree. C., a pressure of 2 to 4
megaPascals (MPa) and hydrogen to feedstock (H.sub.2/feedstock)
molar ratio ranging from 3 to 6.
US2006/0287564 describes a process for increasing the production of
benzene from a hydrocarbon mixture including separating a
hydrocarbon feedstock into a C.sub.6 or lower hydrocarbon stream
and a C.sub.7 or higher hydrocarbon stream. The C.sub.6 or lower
hydrocarbon stream is separated into a non-aromatic hydrocarbon
stream and an aromatic hydrocarbon stream through a solvent
extraction process. The C.sub.7 or higher hydrocarbon stream is
subjected to a reaction in the presence of a catalyst comprising
platinum/tin or platinum/lead.
U.S. Pat. No. 3,957,621 describes a process for processing heavy
reformates from which benzene and lighter components have been
largely removed. The removed stream includes the major portion of
the benzene in the charge and can include a substantial portion of
the toluene.
WO2013/182534 discloses a process for producing BTX from a
C.sub.5-C.sub.12 hydrocarbon mixture using a
hydrocracking/hydrodesulphurisation catalyst. According to
WO2013/182534, the process results in a mixture comprising
substantially no co-boilers of BTX, thus chemical grade BTX can
easily be obtained. WO2015/189058 discloses a further improvement
by selectively recycling back the toluene from the product stream
to be included in the feed.
Often, feed streams are purified before being hydrocracked. For
example, the feed stream may have been desulfurized, depentanised,
and/or processed via extractive distillation. Additionally,
different feed streams comprise different concentrations of
impurities and if the impurity concentration is low, the existence
or issues with that material are not considered or addressed.
Therefore, the various patents discussing hydrocracking do not
address issues in streams that had higher concentrations of some
materials. For example, while WO2013/182534 and WO2015/189058
advantageously provide a chemical grade BTX from many types of feed
streams, they do not mention an amount of difficult co-boilers
(such as 2,4-dimethylpentane and 2,2,3-trimethylbutane) in the feed
stream. The present inventors found that there are some benzene
co-boiler compounds which are more difficult to hydrocrack. As a
result, there are certain benzene co-boilers (such as
2,4-dimethylpentane and 2,2,3-trimethylbutane) that are avoided in
the feed stream and hence in the product stream. In other words,
merely confirming a low amount of co-boilers in the product stream
does not suggest the existence of 2,4-dimethylpentane and
2,2,3-trimethylbutane in the feed stream. Actually, there is a
demand for a hydrocracking process which allows hydrocracking of
substantially all co-boilers of benzene, including
2,4-dimethylpentane and 2,2,3-trimethylbutane. In other words,
there is need for a process that can crack 2,4-dimethylpentane and
2,2,3-trimethylbutane, and produce chemical grade benzene.
It is an object of the present invention to provide a hydrocracking
process in which above and/or other needs are met.
Accordingly, the present invention provides a hydrocracking process
comprising: contacting a hydrocracking feed stream in the presence
of hydrogen with a hydrocracking catalyst under process conditions
including a temperature of 425-580.degree. C., a pressure of
300-5,000 kiloPascals (kPa) gauge and a Weight Hourly Space
Velocity (WHSV) of 0.1-30 h.sup.-1 to produce a hydrocracking
product stream comprising benzene; wherein the hydrocracking feed
stream comprises C.sub.5-C.sub.12 hydrocarbons which comprise
2,4-dimethylpentane and/or 2,2,3-trimethylbutane, wherein the total
amount of 2,4-dimethylpentane and 2,2,3-trimethylbutane is at least
0.5 wt % of the hydrocracking feed stream; wherein the
hydrocracking catalyst comprises a hydrogenation metal in an amount
of 0.010-0.30 wt % with respect to the total catalyst; and wherein
the hydrocracking catalyst comprises a medium pore zeolite having a
pore size of 5-6 .ANG. and a silica to alumina molar ratio of
20-75; preferably the hydrocracking catalyst comprises a medium
pore zeolite having a pore size of 5-6 .ANG. and a silica to
alumina molar ratio of 20-75 and a large pore zeolite having a pore
size of 6-8 .ANG. and a silica to alumina molar ratio of 10-80,
wherein the hydrogenation metal is deposited on the medium pore
zeolite and the large pore zeolite.
It was surprisingly found that certain branched alkanes having
boiling points close to the boiling points of benzene such as
between 75.degree. C. and 90.degree. C., namely 2,4-dimethylpentane
(24DMP) and 2,2,3-trimethylbutane (223TMB), cannot be effectively
hydrocracked by known hydrocracking catalysts. In other words,
either the conversion is less than 90% and/or the deactivation rate
of the catalyst is greater than |-2.5.times.10.sup.-4 hr.sup.-1|.
These branched alkanes which are difficult to hydrocrack are not
present in large amounts in typical feed streams and therefore this
problem is not recognized when using these feed streams. The
recognition of the problem was possible by the present inventors by
hydrocracking specific feed streams with high amounts of 24DMP and
223TMB for the purpose of generating high purity benzene and
analyzing the obtained product stream. The inventors recognized
this problem and found that specific hydrocracking catalyst
according to the invention can solve this problem by effectively
hydrocracking all benzene co-boilers including these benzene
co-boilers. The absence of co-boilers of benzene in the product
stream allows obtaining chemical grade benzene by simple
distillation of the product stream.
It was surprisingly found that the improved conversion of these
benzene co-boilers can be achieved by a catalyst comprising a
hydrogenating metal and a medium pore zeolite having a
silica-to-alumina ratio of 20 to 75. The hydrogenating metal could
be present in greater than or equal to 0.09 wt %, preferably
greater than or equal to 0.10 wt %, e.g., 0.09 to 3 wt %, or 0.10
to 2.5 wt %, based upon a total weight of the catalyst.
It was surprisingly found that the improved conversion of these
benzene co-boilers can be achieved by a catalyst comprising a
medium pore zeolite having a silica to alumina molar ratio of 20-75
and a hydrogenation metal deposited on the zeolite (e.g., a
catalyst comprising a shaped body comprising a medium pore zeolite
and a binder and a hydrogenation metal deposited on the shaped
body), or by a catalyst comprising a medium pore zeolite and a
large pore zeolite and a hydrogenation metal deposited on the
medium pore zeolite and the large pore zeolite.
Accordingly, in some embodiments, the hydrogenation metal is
deposited on a shaped body comprising the medium pore zeolite
having a pore size of 5-6 .ANG. and a silica (SiO.sub.2) to alumina
(Al.sub.2O.sub.3) molar ratio of 20-75 and a binder. It was found
that this leads to a higher conversion of these benzene co-boilers
compared to a similar catalyst where a similar amount of the
hydrogenation metal is deposited on the medium pore zeolite without
a binder. The zeolite in the catalyst may consist of the medium
pore zeolite, or may further comprise a large pore zeolite having a
pore size of 6-8 .ANG. and a silica (SiO.sub.2) to alumina
(Al.sub.2O.sub.3) molar ratio of 10-80.
When the hydrocracking catalyst comprises the medium pore zeolite
having a pore size of 5-6 .ANG. and a silica to alumina molar ratio
of 20-75 and also a large pore zeolite catalyst having a pore size
of 6-8 .ANG. and a silica (SiO.sub.2) to alumina (Al.sub.2O.sub.3)
molar ratio of 10-80 in addition to a medium pore zeolite catalyst,
the hydrocracking catalyst may be in the form of powder and free
from a binder. It was found that pore size (that allows the
molecules to enter the zeolite pores to reach active sites), and
the zeolite silica to alumina ratio (number of acid sites), and
with hydrogenation metal deposited on the large pore zeolite leads
to a higher conversion of the branched alkanes compared to a
similar catalyst where a similar amount of the hydrogenation metal
is deposited only on the medium pore zeolite.
As used herein, the term "C.sub.n hydrocarbons", wherein "n" is a
positive integer, is meant to describe all hydrocarbons having n
carbon atoms. Moreover, the term "C.sub.n+ hydrocarbons" is meant
to describe all hydrocarbon molecules having n or more carbon
atoms. Accordingly, the term "C.sub.5+ hydrocarbons" is meant to
describe a mixture of hydrocarbons having 5 or more carbon
atoms.
Hydrocracking Feed Stream
The hydrocracking feed stream comprises C.sub.5-C.sub.12
hydrocarbons including at least 0.5 wt % of 2,4-dimethylpentane and
2,2,3-trimethylbutane, based upon the total weight of the
hydrocracking feed stream. The hydrocracking feed stream can
comprise C.sub.5-C.sub.12 hydrocarbons including greater than or
equal to 20 wt % (preferably greater than or equal to 30 wt %, or
greater than or equal to 40 wt %) C.sub.6+ non-aromatics, wherein
the C.sub.6+ non-aromatics comprise at least 0.5 wt % of
2,4-dimethylpentane and 2,2,3-trimethylbutane, based upon the total
weight of the hydrocracking feed stream.
The hydrocracking feed stream is a mixture comprising
C.sub.5-C.sub.12 hydrocarbons, preferably having a boiling point
(b.p.) in the range of 30-195.degree. C. Preferably, the
hydrocracking feed stream mainly comprises C.sub.6-C.sub.8
hydrocarbons. The hydrocracking feed stream comprises
2,4-dimethylpentane (b.p. 80.degree. C.) and 2,2,3-trimethylbutane
(b.p. 81.degree. C.). The hydrocracking feed stream may further
comprise other branched alkanes having boiling points between
75.degree. C. and 90.degree. C. These alkanes include
2,2-dimethylpentane (b.p. 78.degree. C.), 3,3-dimethylpentane (b.p.
86.degree. C.), 2,3-dimethylpentane (b.p. 89.degree. C.) and
2-methylhexane (b.p. 90.degree. C.).
The amount of the branched alkanes having boiling points between
75.degree. C. and 90.degree. C. (i.e., benzene co-boiler branched
alkanes) in the hydrocracking feed stream may be at least 0.5 wt %,
at least 1.0 wt %, at least 2.0 wt %, or at least 5.0 wt %, with
respect to total hydrocarbon feed. The amount of the branched
alkanes having boiling points between 75.degree. C. and 90.degree.
C. may be at most 15 wt %, or at most 10 wt %.
The total amount of 2,4-dimethylpentane and 2,2,3-trimethylbutane
in the hydrocracking feed stream may be at least 0.5 wt %, at least
1.0 wt %, at least 2.0 wt %, or at least 5.0 wt %, with respect to
total hydrocarbon feed. The total amount of 2,4-dimethylpentane and
2,2,3-trimethylbutane may be at most 15 wt %, or at most 10 wt %,
with respect to total hydrocarbon feed.
The amount of 2,4-dimethylpentane in the hydrocracking feed stream
may be at least 0.5 wt %, at least 1.0 wt %, at least 2.0 wt %, or
at least 5.0 wt %, with respect to total hydrocarbon feed. The
amount of 2,4-dimethylpentane may be at most 15 wt %, or at most 10
wt %, with respect to total hydrocarbon feed.
The amount of 2,2,3-trimethylbutane in the hydrocracking feed
stream may be at least 0.5 wt %, at least 1.0 wt %, at least 2.0 wt
%, or at least 5.0 wt %, with respect to total hydrocarbon feed.
The amount of 2,2,3-trimethylbutane may be at most 15 wt %, or at
most 10 wt %, with respect to total hydrocarbon feed.
Some types of feed streams may include such high amounts of benzene
co-boiler branched alkanes (e.g., 2,4-dimethylpentane and/or
2,2,3-trimethylbutane). Naphtha derived from natural gas condensate
(e.g., Saudi A-180, Texas shale gas condensate, etc.), pygas
naphtha (benzene-rich naphtha derived from the liquid byproduct of
a steam cracker), straight run naphtha from distillation of crude
oil, naphtha from cracking processes (e.g. FCC, hydrocracking), and
raffinate from reformate, may have various compositions, and some
may have such high amounts of benzene co-boiler branched alkanes
(e.g., 2,4-dimethylpentane and/or 2,2,3-trimethylbutane). It is
noted that only certain types of these feed streams have such high
amounts of benzene co-boiler branched alkanes, such as
2,4-dimethylpentane and/or 2,2,3-trimethylbutane.
In some embodiments, the feed stream used in the process of the
present invention has been depentanised. Preferably, the feed
stream comprises at most 5 wt % of C.sub.5 hydrocarbons, more
preferably at most 4 wt %, at most 3 wt %, at most 2 wt %, at most
1 wt %, or C.sub.5 hydrocarbons.
Preferably, the hydrocracking feed stream is provided by a process
which does not involve the step of removing benzene or removing
C.sub.6 hydrocarbons. This means that intentional removal of
benzene has not been performed in providing the hydrocracking feed
stream or the fresh feed stream. The step of removing benzene
typically induces the removal of co-boilers of benzene. According
to the present invention, the benzene co-boilers present in the
hydrocracking feed stream are advantageously converted to useful
LPG.
Preferably, the hydrocracking feed stream may comprise at least 5
wt % of benzene, for example at least 10 wt % of benzene, at least
20 wt % of benzene, at least 30 wt % of benzene or at least 40 wt %
of benzene, and/or at most 90 wt % of benzene, for example at most
80 wt %, at most 70 wt %, at most 60 wt % or at most 50 wt % of
benzene.
The hydrocracking feed stream contains C.sub.5-C.sub.12
hydrocarbons. For example, the hydrocracking feed stream contains
aromatics and nonaromatics. The aromatics include at least one of
benzene, toluene, and xylene. The aromatics can be present in an
amount of greater than or equal to 40 wt %, e.g., greater than or
equal to 50 wt %, or greater than or equal to 60 wt %, or even
greater than or equal to 70 wt %, based upon a total weight of the
feed stream. The hydrocracking feed stream is contacted in the
presence of hydrogen in a hydrocracking reactor with the
hydrocracking catalyst of the invention.
According to the process of the present invention, a hydrocracking
feed stream comprising a relatively large amount of certain
branched alkanes which are difficult to hydrocrack is efficiently
converted into a mixture comprising substantially no co-boilers of
benzene (e.g., less than or equal to 0.2 wt %). As a result
thereof, chemical grade BTX or chemical grade benzene is obtained
by relatively simple separation methods such as gas-liquid
separation or distillation. The product produced by the
hydrocracking step of the process of the present invention
(hydrocracking product stream) comprises LPG, BTX and methane.
The term "LPG" as used herein refers to the well-established
acronym for the term "liquefied petroleum gas". LPG generally
consists of a blend of C.sub.2-C.sub.4 hydrocarbons i.e. a mixture
of C.sub.2, C.sub.3, and C.sub.4 hydrocarbons.
The term "BTX" as used herein is well known in the art and relates
to a mixture of benzene, toluene and xylenes.
As used herein, the term "chemical grade BTX" relates to a
hydrocarbon mixture comprising less than 5 wt % hydrocarbons other
than benzene, toluene and xylenes, preferably less than 4 wt %
hydrocarbons other than benzene, toluene and xylenes, more
preferably less than 3 wt % hydrocarbons other than benzene,
toluene and xylenes, and most preferably less than 2.5 wt %
hydrocarbons other than benzene, toluene and xylenes.
Furthermore, the "chemical grade BTX" produced by the process of
the present invention comprises less than 1 wt % non-aromatic
C.sub.6+ hydrocarbons, preferably less than 0.7 wt % non-aromatic
C.sub.6+ hydrocarbons, more preferably less than 0.5 wt %
non-aromatic C.sub.6+ hydrocarbons and most preferably less than
0.2 wt % non-aromatic C.sub.6+ hydrocarbons.
As used herein, the term "chemical grade benzene" relates to a
hydrocarbon stream comprising less than or equal to 0.2 wt %
hydrocarbons other than benzene.
The term "aromatic hydrocarbon" is very well known in the art.
Accordingly, the term "aromatic hydrocarbon" relates to cyclically
conjugated hydrocarbon with a stability (due to delocalization)
that is significantly greater than that of a hypothetical localized
structure (e.g. Kekule structure). The most common method for
determining aromaticity of a given hydrocarbon is the observation
of diatropicity in the 1H NMR spectrum, for example the presence of
chemical shifts in the range of from 7.2 to 7.3 ppm for benzene
ring protons.
The hydrocracking product stream produced in the process of the
present invention preferably comprises less than 5 wt % of methane.
Preferably, the hydrocracking product stream produced in the
process of the present invention comprises less than 4 wt % of
methane, more preferably less than 3 wt % methane, even more
preferably less than 2 wt % methane, even more preferably less than
1.5 wt % methane, even more preferably less than 1.4 wt % methane,
even more preferably less than 1.3 wt % methane, even more
preferably less than 1.2 wt % methane, even more preferably less
than 1.1 wt % methane, and most preferably less than 1 wt %
methane.
Preferably, the hydrocracking product stream is also substantially
free from C.sub.5 hydrocarbons. As meant herein, the term
"hydrocracking product stream substantially free from C.sub.5
hydrocarbons" means that said hydrocracking product stream
comprises less than 1 wt % C.sub.5 hydrocarbons, preferably less
than 0.7 wt % C.sub.5 hydrocarbons, more preferably less than 0.6
wt % C.sub.5 hydrocarbons and most preferably less than 0.5 wt %
C.sub.5 hydrocarbons.
Process Conditions
The process conditions under which the hydrocracking of the feed
stream is performed are an important determinant for the
composition of the hydrocracking product stream.
In general, when the space velocity is too high, not all co-boilers
of BTX are hydrocracked, so it will not be possible to obtain a
chemical grade BTX by simple distillation of the product stream.
However, at too low space velocity the yield of methane rises at
the expense of ethane, propane and butane. Also, a higher space
velocity requires smaller reactor volumes and thus a lower CAPEX.
Hence, it is advantageous to perform the process of the invention
at a high space velocity at which substantially all co-coilers of
BTX are hydrocracked.
It was found that the hydrocracking step (b) can advantageously be
performed at a high space velocity while allowing substantially all
co-boilers of BTX to be hydrocracked, due to the high activity of
the catalyst.
Accordingly, in some preferred embodiments, the step (b) is
performed at a Weight Hourly Space Velocity (WHSV) of 0.1-30 per
hour (hr.sup.-1), for example at least 1 hr.sup.-1, at least 2
hr.sup.-1, at least 3 hr.sup.-1, at least 5 hr.sup.-1, at least 6
hr.sup.-1, at least 7 hr.sup.-1 or at least 8 hr.sup.-1, and/or at
most 25 hr.sup.-1, at most 20 hr.sup.-1, at most 15 hr.sup.-1, at
most 10 hr.sup.-1, or at most 9 hr.sup.-1. High WHSV such as at
least 8 hr.sup.-1 allows particularly small reactor volumes and
lower capital expenditure (CAPEX).
It has also been found that step (b) can be operated at a
relatively low temperature. This allows for greater operational
flexibility as well as lower heat duty and may allow longer cycle
lengths. Accordingly, in some preferred embodiments, the step (b)
is performed at a temperature of 425-445.degree. C. In other
embodiments, the step (b) is performed at a temperature of
450-580.degree. C. The higher temperature range results in a high
hydrocracking conversion rate.
The hydrocracking of the feed stream is performed at a pressure of
300-5,000 kPa gauge, more preferably at a pressure of 600-3,000 kPa
gauge, particularly preferably at a pressure of 1,000-2,000 kPa
gauge and most preferably at a pressure of 1200-1600 kPa gauge. By
increasing reactor pressure, conversion of C.sub.5+ non-aromatics
can be increased, but higher pressure also increases the yield of
methane and the hydrogenation of aromatic rings to cyclohexane
species which can be cracked to LPG species. This results in a
reduction in aromatic yield as the pressure is increased and, as
some cyclohexane and its isomer methylcyclopentane are not fully
hydrocracked, a pressure of 1,200-1,600 kPa may result in a high
purity of the resultant benzene.
The hydrocracking step is performed in the presence of an excess of
hydrogen in the reaction mixture. This means that a more than
stoichiometric amount of hydrogen is present in the reaction
mixture that is subjected to hydrocracking. Preferably, the molar
ratio of hydrogen to hydrocarbon species (H.sub.2/HC molar ratio)
in the reactor feed is between 1:1 and 4:1, preferably between 1:1
and 3:1 and most preferably between 2:1 and 3:1. A higher benzene
purity in the product stream can be obtained by selecting a
relatively low H.sub.2/HC molar ratio. In this context the term
"hydrocarbon species" means all hydrocarbon molecules present in
the reactor feed such as benzene, toluene, hexane, cyclohexane,
etc. The composition of the feed and/or the volumetric flow of the
hydrocarbon stream as a vapor, are used to calculate the average
molecular weight of this stream to be able to calculate the correct
hydrogen feed rate. The excess amount of hydrogen in the reaction
mixture suppresses the coke formation which is believed to lead to
catalyst deactivation.
Catalyst
The hydrocracking catalyst used in the process of the present
invention comprises a hydrogenation metal. The hydrocracking
catalyst further comprises a medium pore zeolite having a pore size
of 5-6 Angstroms (.ANG.) and a silica (SiO.sub.2) to alumina
(Al.sub.2O.sub.3) molar ratio of 20 to 75.
In some embodiments, the hydrogenation metal is deposited on a
shaped body comprising the medium pore zeolite and a binder. The
shaped body may further comprise a large pore zeolite having a pore
size of 6-8 .ANG. and a silica (SiO.sub.2) to alumina
(Al.sub.2O.sub.3) molar ratio of 10 to 80.
Examples of the shaped bodies include, but are not limited to,
spherically or cylindrically shaped pellets, tablets, particles and
extrudates. The shaped body typically has an average diameter of
about 0.1 millimeter (mm) to about 7 mm, typically 1.4 mm to 3.5
mm. The diameter is usually measured by slide caliper. The shaped
body typically has an average length of 3 to 8 mm. The average as
used herein is an arithmetic average. One specific example of the
shaped body is cylindrically shaped extrudates with an average
diameter of about 1.6 mm ( 1/16 inch) with an average length of
extrudates about 3 to 8 mm. In such catalyst, the distance between
the hydrogenation metal and the zeolite acid site is less than that
in a mixed catalyst of a shaped zeolite body and hydrogenation
metal supported on a binder. An example of the latter would be a
mixture of ZSM-5 zeolite extrudates and Pt deposited on shaped
Al.sub.2O.sub.3.
In some embodiments, the hydrocracking catalyst may be in the form
of powder and free from a binder. In this case, the hydrocracking
catalyst further comprises a medium pore zeolite having a pore size
of 5-6 .ANG. and a silica (SiO.sub.2) to alumina (Al.sub.2O.sub.3)
molar ratio of 20-75 and a large pore zeolite having a pore size of
6-8 .ANG. and a silica (SiO.sub.2) to alumina (Al.sub.2O.sub.3)
molar ratio of 10-80. The hydrogenation metal and optionally La
and/or Ga are deposited on the medium pore zeolite and the large
pore zeolite.
Zeolite
The hydrocracking catalyst used according to the invention
comprises a medium pore zeolite having a pore size of 5-6 .ANG..
The hydrocracking catalyst used according to the invention may
further comprise a large pore zeolite having a pore size of 6-8
.ANG.. The terms "medium pore zeolite" and "large pore zeolite" are
commonly used in the field of zeolite catalysts.
Zeolites are well-known molecular sieves having three dimensional
structures with well-defined channels, pores, cavities with defined
pore size. As used herein, the term "zeolite" or "aluminosilicate
zeolite" relates to an aluminosilicate molecular sieve. An overview
of their characteristics is for example provided by the chapter on
Molecular Sieves in Kirk-Othmer Encyclopedia of Chemical
Technology, Volume 16, p 811-853; in Atlas of Zeolite Framework
Types, 5th edition, (Elsevier, 2001). In this Atlas of Zeolite
Framework Types, various zeolites are listed based on ring
structure. Zeolites of the 8-ring structure type are called small
pore zeolites.
Preferably, the zeolite present in the catalyst is in a hydrogen
form or a NH.sub.4-form, i.e. having at least a portion of the
original cations associated therewith replaced by H.sup.+ ions or
NH.sub.4.sup.+ ions, respectively. Various methods to convert an
aluminosilicate zeolite to the hydrogen form can be used. A first
method involves direct treatment employing an acid, for example a
mineral acid (HNO.sub.3, HCl, etc.). A second method involves
direct exchange using an ammonium salt (e.g. NH.sub.4NO.sub.3)
followed by calcination. The zeolite can optionally contain up to
trace levels of other cations such as Na (wherein a trace level is
at most 0.05 wt % based upon the total weight of the zeolite).
Possible zeolites include, but are not limited to, ZSM-5, MCM-22,
ZSM-11, beta zeolite, EU-1 zeolite, faujasite (zeolite Y),
ferrierite and mordenite.
Possible medium pore zeolites are 10-ring zeolites, i.e. the pore
is formed by a ring consisting of 10 tetrahedra of [SiO.sub.4] and
[AlO.sub.4].sup.-. The negative charge arising from
[AlO.sub.4].sup.- is neutralized by a cation in the zeolite.
Preferably, the medium pore zeolite is a ZSM-5 zeolite.
The silica (SiO.sub.2) to alumina (Al.sub.2O.sub.3) molar ratio of
the medium pore zeolite is in the range of 20-75. This shows the
optimum catalyst hydrocracking performance to obtain desired
benzene purity from a feedstock containing benzene co-boilers.
Means and methods for quantifying the SiO.sub.2 to Al.sub.2O.sub.3
molar ratio of a zeolite include, but are not limited to AAS
(Atomic Absorption Spectrometer), ICP (Inductively Coupled Plasma
Spectrometry) analysis, and XRF (X-ray fluorescence). It is noted
that the SiO.sub.2 to Al.sub.2O.sub.3 molar ratio referred herein
is meant as the ratio in the zeolite prior to being mixed with the
other components. Preferably, the SiO.sub.2 to Al.sub.2O.sub.3
molar ratio is measured by XRF.
Preferably, the silica to alumina molar ratio of the medium pore
zeolite is in the range of 20-50. At such ratio, the hydrocracking
of the benzene co-boilers is especially efficient. Even more
preferably, the silica to alumina molar ratio of the medium pore
zeolite is in the range of 20-30 or 21-29. At such ratio, the
resistance of the catalyst to catalyst deactivation is high.
In some embodiments, the zeolite present in the catalyst consists
of the medium pore zeolite. In other embodiments, the hydrocracking
catalyst is a combination of the medium pore zeolite and the large
pore zeolite.
Suitable large pore zeolites are 12-ring zeolites. Preferably, the
large pore zeolite includes at least one of Y zeolite, beta
zeolite, and mordenite.
The silica (SiO.sub.2) to alumina (Al.sub.2O.sub.3) molar ratio of
the large pore zeolite is in the range of 5 to 100, for example 10
to 80, preferably 10 to 60.
Preferably, the zeolite in the hydrocracking catalyst comprises
70-100 wt % of the medium pore zeolite and 0-30 wt % of the large
pore zeolite, for example 75-95 wt % of the medium pore zeolite and
5-25 wt % of the large pore zeolite, or 75-85 wt % of the medium
pore zeolite and 15-25 wt % of the large pore zeolite, with respect
to the total amount of the zeolite.
Hydrogenation Metal
Preferably, the hydrogenation metal is at least one element
selected from Group 10 of the periodic table of Elements, rhodium,
and iridium. The preferred Group 10 elements are palladium and
platinum, particularly platinum. In other words, the hydrogenation
metal can consist of platinum. Optionally, the hydrocracking
catalyst can be free of metals other than the Group 10 metals of
the Periodic Table of Elements, rhodium, and iridium; preferably
free of metals other than palladium and platinum. As used herein
"free of metals" means that no other metals were added to the
hydrocracking catalyst.
Binder
In some embodiments, the catalyst comprises a binder. The binder
material can be an inorganic oxide material. The binder material
can comprise an aluminum-containing or silicon-containing material
such as silica, alumina, clay, aluminum phosphate, silica-alumina,
or combinations comprising at least one of the foregoing. Alumina
(Al.sub.2O.sub.3) is the preferred binder. The catalyst can
comprise up to 99 wt %, e.g., 1 to 99 wt %, for example 10 to 90 wt
%, 10 to 50 wt % or 20 to 40 wt % of the binder based on the total
weight of the catalyst.
Preferably, the binder has been treated with a mineral acid such as
nitric acid, hydrochloric acid, phosphoric acid, or sulfuric acid,
preferably nitric acid. Treating the binder with a mineral acid
improves physical strength of the formed catalyst.
It will be appreciated that materials such as SiC used for diluting
the catalyst before loading to the reactor are not considered as a
binder and is not part of the catalyst.
Metal Amount
The catalyst according to the process of the present invention
comprises 0.010-0.30 wt. % of the hydrogenating metal. In the
context of the present invention, the term "wt %" when relating to
the metal content as comprised in a catalyst relates to the wt % of
said metal in relation to the total weight of the catalyst. The
amount of the hydrogenating metal in the catalyst can be determined
e.g. by subjecting the catalyst to XRF or ICP.
Even more preferably, the catalyst comprises at least 0.030 wt %,
at least 0.050 wt %, at least 0.075 wt %, at least 0.10 wt %, at
least 0.125 wt % or at least 0.20 wt %, of the hydrogenating metal
in relation to the total weight of the catalyst. At such amount,
the resistance of the catalyst to catalyst deactivation is high.
The catalyst may comprise at most 0.275 wt % of the hydrogenating
metal in relation to the total weight of the catalyst.
The hydrocracking catalyst may further comprise La and/or Ga. The
total amount of La and/or Ga may be 0.10-0.40 wt % of the total
weight of the catalyst. However, in some embodiments, the
hydrocracking catalyst does not comprise La and Ga, i.e. comprises
less than 0.01 parts Ga and less than 0.01 parts, preferably
comprises less than 0.005 parts Ga and less than 0.005 parts (on
the basis of 100 parts by weight of total catalyst).
Desirably, the hydrocracking catalyst has a sufficient
hydrogenation activity to hydrogenate unsaturated non-aromatic
hydrocarbons. Accordingly, it is preferred that the catalyst does
not comprise secondary metals, such as tin, lead, or bismuth that
inhibit the hydrogenation activity of the hydrogenation metal.
Preferably, the hydrocracking catalyst comprises less than 0.01
parts tin and less than 0.02 parts lead and less than 0.01 parts
bismuth (on the basis of 100 parts by weight of the total
catalyst), preferably less than 0.005 parts tin and less than 0.01
parts lead and less than 0.005 parts bismuth (on the basis of 100
parts by weight of total catalyst).
Further, preferably, the hydrocracking catalyst can comprise less
than 0.01 parts molybdenum (on the basis of 100 parts by weight of
the total catalyst).
Process for Preparation of Catalyst
The catalyst may be made by depositing the hydrogenation metal (and
optionally La and/or Ga) on the zeolite for preparing a catalyst in
powder form. The catalyst may also be made by depositing the
hydrogenation metal (and optionally La and/or Ga) on the shaped
body comprising the zeolite and the binder, e.g. by a wet or vapor
phase impregnation or by an ion-exchange method. Examples of the
preparation method for the catalyst wherein the hydrogenation metal
is Pt uses (NH.sub.3).sub.4Pt(NO.sub.3).sub.2,
(NH.sub.3).sub.4PtCl.sub.2 or (NH.sub.3).sub.4Pt(OH).sub.2 as a
platinum source usually in combination with NH.sub.4C.sub.1.
Another example of the preparation method for the catalyst wherein
the hydrogenation metal is Pt uses H.sub.2PtCl.sub.6 as a platinum
source. The method wherein H.sub.2PtCl.sub.6 is used as the
platinum source may be preferable in that NH.sub.4Cl is not
needed.
The feed stream is cracked by the catalyst to produce a
hydrocracking product stream comprising LPG and BTX. The process
may further comprise separating BTX or benzene from the
hydrocracking product stream. The hydrocracking product stream may
be subjected to separation by standard means and methods suitable
for separating methane and unreacted hydrogen comprised in the
hydrocracking product stream as a first separate stream, the LPG
comprised in the hydrocracking product stream as a second separate
stream and BTX as a third separate stream. Preferably, the stream
comprising BTX is separated from the hydrocracking product stream
by gas-liquid separation or distillation. Benzene may be further
separated from the stream comprising BTX.
One non-limiting example of such a separation method of the
hydrocracking product stream includes a series of distillation
steps. The first distillation step at moderate temperature is to
separate most of the aromatic species (liquid product) from the
hydrogen, H.sub.2S, methane and LPG species. The gaseous stream
from this distillation is further cooled (to about -30.degree. C.)
and distilled again to separate the remaining aromatic species and
most of the propane and butane. The gaseous product (mainly
hydrogen, H.sub.2S, methane, and ethane) is then further cooled (to
about -100.degree. C.) to separate the ethane and leave the
hydrogen, H.sub.2S and methane in the gaseous stream that will be
recycled back to the hydrocracking reactor. To control the levels
of H.sub.2S and methane in the reactor feed, a proportion of this
recycle gas stream is removed from the system as a purge. The
quantity of material that is purged depends on the levels of
methane and H.sub.2S in the recycle stream which in-turn depend on
the feed composition. As the purge will contain mainly hydrogen and
methane it is suitable for use as a fuel gas or may be further
treated (e.g. via a pressure swing adsorption unit) to separately
recover a high purity hydrogen stream and a methane/H.sub.2S stream
which can be used as a fuel gas.
In a further embodiment, the present invention relates to a process
for producing benzene from a feed stream comprising
C.sub.5-C.sub.12 hydrocarbons, wherein the said process comprises
the hydrocracking process further comprising separating BTX or
benzene from the hydrocracking product stream, further comprising
the step of contacting BTX (or only the toluene and xylenes
fraction of said BTX produced) with hydrogen under conditions
suitable to produce a hydrodealkylation product stream comprising
benzene and fuel gas.
The conditions suitable to produce a hydrodealkylation product
stream comprising benzene and fuel gas are well-known and are
described in detail e.g. in WO2013/182534, incorporated herein by
reference.
Processes for hydrodealkylation of hydrocarbon mixtures comprising
C.sub.6-C.sub.9 aromatic hydrocarbons include thermal
hydrodealkylation and catalytic hydrodealkylation; see e.g. WO
2010/102712 A2. Catalytic hydrodealkylation is preferred as this
hydrodealkylation process generally has a higher selectivity
towards benzene than thermal hydrodealkylation. Preferably
catalytic hydrodealkylation is employed, wherein the
hydrodealkylation catalyst is selected from supported chromium
oxide catalyst, supported molybdenum oxide catalyst, platinum on
silica or alumina and platinum oxide on silica or alumina.
The process conditions useful for hydrodealkylation, also described
herein as "hydrodealkylation conditions", can be easily determined
by the person skilled in the art. The process conditions used for
thermal hydrodealkylation are for instance described in DE 1668719
A1 and include a temperature of 600-800.degree. C., a pressure of
3-10 MPa gauge and a reaction time of 15-45 seconds. The process
conditions used for the preferred catalytic hydrodealkylation
preferably include a temperature of 500-650.degree. C., a pressure
of 3.5-7 MPa gauge and a Weight Hourly Space Velocity of 0.5-2
h.sup.-1; see also Handbook of Commercial Catalysts: Heterogeneous
Catalysts ed. Howard F. Rase (2000) Loc. cit.
The hydrodealkylation product stream can be separated into a liquid
stream (containing benzene and other aromatics species) and a gas
stream (containing hydrogen, H.sub.2S, methane, and other low
boiling point hydrocarbons) by a combination of cooling and
distillation. The liquid stream may be further separated, by
distillation, into a benzene stream, a C.sub.7 to C.sub.9 aromatics
stream and a heavy aromatic stream. The C.sub.7 to C.sub.9 aromatic
stream, or some part of it, may be fed back to reactor section as a
recycle to increase overall conversion and benzene yield. The heavy
aromatic stream, which contains polyaromatic species such as
biphenyl, is preferably not recycled to the reactor but may be
exported as a separate product stream. The gas stream contains
significant quantities of hydrogen and may be recycled back, via a
recycle gas compressor, to the reactor section. A recycle gas purge
may be used to control the concentrations of methane and H.sub.2S
in the reactor feed.
The invention is now elucidated by way of the following examples,
without however being limited thereto.
EXAMPLES
Below table shows the analyzed composition of a pygas (Platfiner)
stream. It can be understood that this pygas comprises essentially
no amount of hydrocarbons having a boiling point close to benzene
except for cyclohexane, methylcyclopentane (MCP) and
1,3-dimethylcyclopentane (1,3-DMCP). It comprises no detectable
amount of 2,4-dimethylpentane (BP 80.degree. C.) or
2,2,3-trimethylbutane (BP 81.degree. C.).
Pygas Composition (Platfiner)
TABLE-US-00001 Component C Number BP, .degree. C. wt % butane 4 -1
0.02 methylbutane 5 27.7 0.34 pentane 5 36.1 0.39 2-methylpentane 6
60 3.93 3 methylpentane 6 63 2.58 hexane 6 68 7.59
methylcyclopentane 6 72 7.98 benzene 6 80.1 47.87 cyclohexane 6
80.7 3.24 trans 1,3-dimethylcyclopentane 7 91 0.47
1,3-dimethylcyclopentane 7 92 1.15 2,2,4-trimethylpentane 7 98 0.37
heptane 7 98 0.80 methylcyclohexane 7 101 0.57 ethylcyclopentane 7
103 0.66 toluene 7 111 13.59 octane 8 126 0.05 ethylbenzene 8 136
3.07 m/p-xylene 8 139/140 1.41 o-xylene 8 144 0.44 nonane 9 151
0.05 iso propyl benzene 9 152 0.04 propyl benzene 9 159 0.01
1-methyl-(3&4)-ethylbenzene 9 152 0.06 1,3,5-trimethylbenzene 9
163-166 0.01 1-methyl 2-ethyl benzene 9 152 0.01 pseudocumene 9
168.5 0.02 indane 9 176.5 0.01 butyl benzene 10 183 0.01
1,3-diethyl benzene 10 182 0.01
Reactor and Catalyst Test Conditions
Referring to Experiments 1 to 5, the hydrocracking of a hydrocarbon
feed stream employing catalysts described in this application were
performed using stainless steel tube reactor as described below.
0.5 grams (g) of catalyst (sized 20-40 mesh) was diluted to 4
milliliters (ml) by premixing with SiC (30 grit) and was loaded in
a reactor.
Reactor description: 3/8 inch ('') inch tube, 0.035'' wall
thickness. 1/8'' thermocouple with a 1/4'' spacer bar;
12''.times.1'' aluminum over-sleeve; reactor bed is approximately 4
inches in length in the center of the sleeve.
The catalyst was pre-activated (drying, Pt reduction) by subjecting
it to 100 standard cubic centimeters (sccm) of H.sub.2 per minute
at 130.degree. C. under 50 pounds per square inch gauge (psig) for
2 hours and subsequently the temperature was raised to 350.degree.
C. (at 50 psig) for reduction under 200 sccm of H.sub.2 (with 50
parts per million by weight (ppm) of H.sub.2S) for 30 min.
Hereinafter, standard feed refers to a feed consisting of 70 wt %
benzene, 15 wt % methylcyclopentane and 15 wt % 3-methylpentane. In
all Experiments 1 to 5, the standard feed was first introduced to
the reactor containing specific catalyst at hydrocracking reaction
conditions as described below and continued for a minimum of 15
hours to establish a steady cracking activity. Subsequently, the
standard feed was replaced by the feed containing a specific
branched hydrocarbon as described for each experiment. All
components of the hydrocracking feed stream are Aldrich regent
grade chemicals dried under 4 A molecular sieves overnight.
The standard feed was introduced to the reactor at a temperature of
470.degree. C. and a pressure of 200 psig. The molar ratio of
H.sub.2 to the hydrocarbons was 4 to 1, and the H.sub.2S content
was 50 ppm based on the total hydrocarbon and H.sub.2 feed. In all
experiments, the same WHSV was maintained.
Experiment 1
Feeds were prepared by adding one of the compounds shown in Table 1
(hydrocarbons having boiling point of 75 to 90.degree. C.) to a
feed containing benzene, methylcyclopentane (MCP) and
3-methylpentane (3MP). The resulting feeds contained 70 wt %
benzene, 15 wt % methylcyclopentane (MCP), 10 wt % 3-methylpentane
(3MP), and 5 wt % of one of the compounds shown in Table 1. In each
of examples (column 1 in Table 1), after initial stable activity
with the standard feed, each of the feeds containing 5 wt % of one
of the components (column 1 in Table 1) was subjected to
hydrocracking at 470.degree. C., about WHSV 10/h and 200 psig.
The feed stream contains H.sub.2 (H.sub.2/HC molar 4) and 50 ppm S
(fed H.sub.2S). The hydrocracking catalyst used was a powder
catalyst of ZSM-5 deposited with Pt (no binder), wherein the amount
of Pt was 0.03 wt % of the total catalyst and the silica to alumina
molar ratio of the ZSM-5 was 50.
The result of the conversion is shown in Table 1. It can be seen
that the conversion of 2,4-dimethylpentane (24DMP) and
2,2,3-trimethylbutane (223TMB) is low, e.g., under 80% and under
30% conversion respectively. It can be understood that it is
difficult to obtain a product stream with substantially no benzene
co-boilers from a feed stream comprising large amounts of 24DMP
and/or 223TMB, while other hydrocarbons can be substantially
completely converted.
TABLE-US-00002 TABLE 1 Nonaromatics Carbon # b.p., .degree. C. %
Conversion 2,2-dimethylpentane 7 78 100.0 2,4-dimethylpentane 7 80
78.3 2,2,3-trimethylbutane 7 81 27.7 2,3-dimethylpentane 7 89 99.7
2-methylhexane 7 90 100 cyclohexane 6 81 99.4
Experiment 2: Effect of Extrudate vs Powder on 223TMB Cracking
In each of the examples, after initial stable activity with the
standard feed, a feed containing 70 wt % benzene, 15 wt %
methylcyclopentane (MCP), 10 wt % 3-methylpentane (3MP) and 5 wt %
of 2,2,3-trimethylbutane (223TMB) was subjected to hydrocracking at
470.degree. C., about WHSV 10/h and 200 psig. The feed stream
contains H.sub.2 (H.sub.2/HC molar 4) and 50 ppm S (fed
H.sub.2S).
In CEx 1-2, the catalysts used were a Pt deposited ZSM-5 powder
catalyst with no binder. In Ex 3-5, the catalysts used were in the
form of an extrudate of ZSM-5 and alumina on which Pt was
deposited. The silica-to-alumina molar ratio of the ZSM-5 is 50.
The amount of Pt in the catalyst is shown in Table 2.
TABLE-US-00003 TABLE 2 223TMB conversion Catalyst Averaged Exam- Pt
Conversion during Deactivation ple Pt on (wt %) (%) tos (h)
rate.sup.1 (hr.sup.-1) CEx1 ZSM-5 0.032 27.7 30-79 not estimated
powder CEx2 ZSM-5 0.102 36.5 56-71 not estimated powder Ex3 ZSM-5
0.067 52.7 30-48 |-4.2 .times. 10.sup.-3| extrudates Ex4 ZSM-5 0.15
74.6 50-60 |-6.1 .times. 10.sup.-3| extrudates Ex5 ZSM-5 0.25 83.7
52-70 |-3.4 .times. 10.sup.-3| extrudates .sup.1Deactivation rate:
absolute value of the decrease of % conversion of 223TMB per hour
calculated during time-on-stream (tos) is indicated.
It can be understood that an increased amount of Pt shows an
increase in the conversion of 223TMB for both the powder form and
the extrudate form. However, the extrudate shows a higher
conversion than the powder even at a low Pt content (comparison of
CEx2 with 0.102 wt % Pt vs Ex3 with 0.067 wt % Pt). It can be
concluded that the extrudate shows a better conversion than the
powder at the same Pt content.
Experiment 3: Effect of Large Pore Zeolite on 24DMP Cracking
In each of the examples, after initial stable activity was achieved
with the standard feed, a feed containing 70 wt % benzene, 15 wt %
methylcyclopentane (MCP), 10 wt % 3-methylpentane (3MP), and 5 wt %
of 24DMP was subjected to hydrocracking at 470.degree. C., about
WHSV 10/h and 200 psig. The feed stream contained H.sub.2
(H.sub.2/HC molar 4) and 50 ppm S (fed H.sub.2S).
In CEx 6-7, the catalysts used were a Pt deposited ZSM-5 powder
catalyst with no binder. In Ex 8-11, the catalysts used were in the
form of a physical mixture of a Pt deposited ZSM-5 powder catalyst
and a Pt deposited large pore zeolite powder catalyst. The
silica-to-alumina molar ratio of the ZSM-5 is 50. The
silica-to-alumina molar ratio of the large pore zeolites are shown
in Table 3. The amounts of Pt in the catalyst are shown in Table
3.
TABLE-US-00004 TABLE 3 24DMP Zeolite Catalyst conversion Pt, wt %,
avgd in mixed % during Deactivation Example Description
zeolite.sup.2 conv tos (h) rate.sup.1 (hr.sup.-1) CEx 6 100 wt % of
Pt(0.03%)/ZSM-5 0.03 78.3 31-82 |-3.3 .times. 10.sup.-4| CEx 7 100
wt % of Pt(0.097%)/ZSM-5 0.097 93.9 29-48 |-4.9 .times. 10.sup.-4|
Ex 8 80 wt % of 20 wt % of 0.051 92.3 18-87 |-2.0 .times.
10.sup.-4| Pt(0.04%)/ Pt(0.094%)/ ZSM-5, SAR 50 HY, SAR 60 Ex 9 80
wt % of 20 wt % of 0.051 93.3 17-45 |-0.96 .times. 10.sup.-4|
Pt(0.04%)/ Pt(0.096%)/ ZSM-5, SAR 50 HY, SAR 30 Ex 10 80 wt % of 20
wt % of 0.054 98.1 200-255 |-1.7 .times. 10.sup.-5| Pt(0.04%)/
Pt(0.11%)/ ZSM-5, SAR 50 Beta, SAR 20 Ex 11 80 wt % of 20 wt % of
0.052 99 17-39 no apparent Pt(0.04%)/ Pt(0.102%)/ deactivation
ZM-5, SAR 50 Mordenite, SAR 20 SAR = silica (SiO.sub.2) to alumina
(Al.sub.2O.sub.3) molar ratio .sup.1Deactivation rate: absolute
value of the decrease of % conversion of 24DMP per hour calculated
during time-on-stream (tos) indicated. .sup.2Estimated from Pt
contents in ZSM-5 and the other large pore zeolite.
The mixtures of a medium pore zeolite catalyst and a large pore
zeolite catalyst show a high conversion rate of more than 92% at Pt
content of 0.051 wt %. From CEx6 and CEx7, a catalyst without a
large pore zeolite having a Pt content of 0.051 wt % can be
estimated to have a conversion of 83.2% from the relationship
between the Pt content and the conversion. It is further noted that
the catalyst containing a mixture of a large pore zeolite catalyst
and a medium pore zeolite catalyst had nearly half the amount of
hydrogenating metal yet attained a conversion of 24DMP of greater
than 90%. It can be concluded that a catalyst containing a mixture
of a large pore zeolite catalyst and a medium pore zeolite catalyst
shows an unexpectedly better conversion than a catalyst containing
only medium pore zeolite catalyst, and the deactivation was slower.
For deactivation rate, the absolute values of the rate are
compared. The larger the absolute value of the deactivation rate,
the faster the catalyst deactivates.
Experiment 4: Effect of Silica-to-Alumina Ratio (SAR) on 24DMP
Cracking
In each of examples, after initial stable activity with the
standard feed, a feed containing 70 wt % benzene, 15 wt %
methylcyclopentane (MCP), 10 wt % 3-methylpentane (3MP), and 5 wt %
of 24DMP was subjected to hydrocracking at 470.degree. C., about
WHSV 10/h and 200 psig. The feed stream contains H.sub.2
(H.sub.2/HC molar 4) and 50 ppm S (fed H.sub.2S). The hydrocracking
catalysts comprised Pt deposited ZSM-5 powder catalyst with no
binder. The amount of Pt and the silica-to-alumina ratio of the
ZSM-5 are summarized in Table 4. The result of the conversion is
shown in Table 4.
TABLE-US-00005 TABLE 4 24DMP Conversion averaged during
Deactivation Example SiO.sub.2/Al.sub.2O.sub.3 Pt (wt %) % conv tos
(h) rate.sup.1 (hr.sup.-1) REx12 80 0.094 62.9 32-100 |-7.6 .times.
10.sup.-4| REx13 50 0.097 94.0 29-48 |-4.9 .times. 10.sup.-4| REx14
30 0.099 98.6 27-43 not observed REx15 23 0.101 99.6 28-100 not
observed .sup.1Deactivation rate: absolute value of the decrease of
% conversion per hour calculated during time-on-stream (tos)
indicated.
The comparison of Reference Experiments 12-15 shows that a lower
molar ratio of silica to alumina led to a higher conversion rate of
24DMP. In particular, the catalysts having a molar ratio of silica
to alumina of 23-30 showed an extremely high conversion and no
decline in 24DMP conversion with no apparent catalyst deactivation
during the time on stream indicated.
Experiment 5: Effect of Pt Loading on 24DMP Cracking
In each of examples, after initial stable activity with the
standard feed, a feed containing 70 wt % benzene, 15 wt %
methylcyclopentane (MCP), 10 wt % 3-methylpentane (3MP), and 5 wt %
of 24DMP was subjected to hydrocracking at 470.degree. C., about
WHSV 10/h and 200 psig. The feed stream contains H.sub.2
(H.sub.2/HC molar 4) and 50 ppm S (fed H.sub.2S). The hydrocracking
catalysts used were extrudates of ZSM-5 and a binder, wherein Pt
was deposited on the extrudates. The silica-to-alumina molar ratio
of the ZSM-5 was 50. The amount of Pt is summarized in Table 5. The
result of the conversion is shown in Table 5.
TABLE-US-00006 TABLE 5 24DMP Conversion averaged Catalyst
conversion during Deactivation Example Pt (wt %) (%) tos (hr)
rate.sup.1 (hr.sup.-1) Ex 16 0.067 90.4 126-169 |-5.1 .times.
10.sup.-4| Ex 17 0.079 93.5 47-71 |-6.0 .times. 10.sup.-4| Ex 18
0.15 99.5 50-62 |-3.5 .times. 10.sup.-4| Ex 19 0.25 99.8 59-70
|-1.7 .times. 10.sup.-4| .sup.1Deactivation rate: decrease of %
conversion per hour calculated during time-on-stream (tos)
indicated.
A higher Pt amount led to a higher conversion of 24DMP. A higher Pt
amount also led to less catalyst deactivation.
Desirably, the hydrocracking catalyst, when cracking
2,4-dimethylpentane and/or 2,2,3-trimethylbutane, has a
deactivation rate (i.e., decline in conversion percent) of less
than the absolute value of -3.5.times.10.sup.-4 hr.sup.-1 (e.g.,
|-3.5.times.10.sup.-4 hr.sup.-1|), for example, less than or equal
to |-3.0.times.10.sup.-4 hr.sup.-1|, or less than or equal to
|-2.5.times.10.sup.-4 hr.sup.-1|, or less than or equal to
|-2.0.times.10.sup.-4 hr.sup.-1|, most preferably less than or
equal to |-1.0.times.10.sup.-4 hr.sup.-1|.
Set forth below are some embodiments of the methods disclosed
herein.
Embodiment 1
A method of hydrocracking at least one of 2,4-dimethylpentane and
2,2,3-trimethylbutane, comprising: contacting a hydrocracking feed
stream in the presence of hydrogen with a hydrocracking catalyst
under process conditions including a temperature of 425-580.degree.
C., a pressure of 300-5000 kPa gauge and a Weight Hourly Space
Velocity of 0.1-30 h.sup.-1 to produce a hydrocracking product
stream comprising benzene (e.g., comprising BTX and LPG); wherein
the hydrocracking feed stream comprising C.sub.5-C.sub.12
hydrocarbons which includes at least 0.5 wt % of
2,4-dimethylpentane and/or 2,2,3-trimethylbutane, based upon a
total weight of the hydrocracking feed stream; and wherein the
hydrocracking catalyst comprises a hydrogenation metal in an amount
of 0.010-0.30 wt % with respect to the total catalyst; and wherein
the hydrocracking catalyst comprises a medium pore zeolite having a
pore size of 5-6 .ANG. and a silica to alumina molar ratio of
20-75; preferably the hydrocracking catalyst comprises a medium
pore zeolite having a pore size of 5-6 .ANG. and a silica to
alumina molar ratio of 20-75 and a large pore zeolite having a pore
size of 6-8 .ANG. and a silica to alumina molar ratio of 10-80,
wherein the hydrogenation metal is deposited on the medium pore
zeolite and the large pore zeolite.
Embodiment 2
The process according to Embodiment 1, wherein the total amount of
2,4-dimethylpentane and 2,2,3-trimethylbutane in the hydrocracking
feed stream is at least 1.0 wt %, at least 2.0 wt %, or at least
5.0 wt %, with respect to total hydrocarbon feed.
Embodiment 3
The process according to any one of Embodiments 1-2, wherein the
hydrocracking catalyst comprises 0.08 to 0.25 wt % hydrogenation
metal, 15 wt % to 25 wt % alumina, and a balance being the medium
pore zeolite.
Embodiment 4
The process according to any one of Embodiments 1-2, wherein the
hydrocracking catalyst is in the form of powder and is free from a
binder.
Embodiment 5
The process according to any one of the preceding embodiments,
wherein the silica to alumina molar ratio of the medium pore
zeolite is in the range of 20-50, preferably 20 to 30, more
preferably 21 to 29.
Embodiment 6
The process according to any one of the preceding embodiments,
wherein a conversion of any 2,4-dimethylpentane is greater than or
equal to 90%, preferably greater than or equal to 95%, or greater
than or equal to 98%; and the conversion of any
2,2,3-trimethylbutane is greater than or equal to 90%, preferably
greater than or equal to 95%, or greater than or equal to 98%.
Embodiment 7
The process according to any one of the preceding embodiments,
wherein the hydrogenating metal is at least one element selected
from Group 10 of the periodic table of elements, rhodium, and
iridium; preferably at least one metal selected from palladium and
platinum; most preferably platinum.
Embodiment 8
The process according to any one of the preceding embodiments,
wherein the hydrocracking catalyst comprises at least 0.030 wt %,
at least 0.050 wt %, at least 0.075 wt %, at least 0.10 wt %, at
least 0.125 wt % or at least 0.20 wt %, of the hydrogenating metal
in relation to the total weight of the catalyst.
Embodiment 9
The process according to any one of the preceding embodiments,
wherein the hydrocracking catalyst comprises La and/or Ga,
preferably at an amount of 0.10-0.40 wt % of the total weight of
the catalyst.
Embodiment 10
The process according to any one of the preceding embodiments,
wherein the zeolite in the hydrocracking catalyst comprises 70-100
wt % of the medium pore zeolite and 0-30 wt % of the large pore
zeolite with respect to the total amount of the zeolite.
Embodiment 11
The process according to any one of the preceding embodiments,
wherein the zeolite in the hydrocracking catalyst comprises 75-95
wt % of the medium pore zeolite and 5-25 wt % of the large pore
zeolite with respect to the total amount of the zeolite.
Embodiment 12
The process according to any one of the preceding embodiments,
wherein the process comprises separating benzene from the
hydrocracking product stream.
Embodiment 13
The process according to any one of the preceding embodiments,
wherein the hydrocracking catalyst has a deactivation rate of less
than |-3.5.times.10.sup.-4 per hour|, preferably less than or equal
to |-3.0.times.10.sup.-4 per hour|, or less than or equal to
|-2.5.times.10.sup.-4 per hour|, most preferably less than or equal
to |-2.0.times.10.sup.-4 per hour|.
Embodiment 14
The process according to any one of the preceding embodiments,
wherein the medium pore zeolite comprises a ZSM-5.
Embodiment 15
The process according to any one of the preceding embodiments,
wherein the large pore zeolite comprises a mordenite.
Embodiment 16
The process according to any one of the preceding embodiments,
wherein the hydrocracking catalyst has a deactivation rate of less
than or equal to |-2.5.times.10.sup.-4 hr.sup.-1|, or less than or
equal to |-2.0.times.10.sup.-4 hr.sup.-1|, or equal to
|-1.0.times.10.sup.-4 hr.sup.-1|.
Embodiment 17
The process according to any one of the preceding embodiments,
wherein a conversion percent of the 2,4-dimethylpentane and
2,2,3-trimethylbutane is greater than or equal to 95%, preferably
greater than or equal to 98%, or greater than or equal to 99%, or
greater than or equal to 99.5%.
Embodiment 18
The process according to any one of the preceding embodiments,
wherein greater than or equal to 95%, preferably greater than or
equal to 98%, more preferably, greater than or equal to 99.8%, the
C.sub.5-C.sub.12 hydrocarbons are cracked.
Embodiment 19
The use of a hydrocracking catalyst to crack at least one of
2,4-dimethylpentane and 2,2,3-trimethylbutane, wherein the
hydrocracking catalyst comprises a medium pore zeolite having a
pore size of 5-6 .ANG. and a silica to alumina molar ratio of
20-75; preferably the hydrocracking catalyst comprises a medium
pore zeolite having a pore size of 5-6 .ANG. and a silica to
alumina molar ratio of 20-75 and a large pore zeolite having a pore
size of 6-8 .ANG. and a silica to alumina molar ratio of 10-80,
wherein the hydrogenation metal is deposited on the medium pore
zeolite and the large pore zeolite.
Embodiment 20
The use according to Embodiment 19, wherein the hydrocracking
catalyst comprises 0.08 to 0.25 wt % hydrogenation metal, 15 wt %
to 25 wt % alumina, and a balance being the medium pore
zeolite.
Embodiment 21
The use according to Embodiment 19, wherein the hydrocracking
catalyst is in the form of powder and is free from a binder.
Embodiment 22
The use according to any one of Embodiments 19-21, wherein the
hydrogenating metal is at least one element selected from Group 10
of the periodic table of elements, rhodium, and iridium; preferably
at least one metal selected from palladium and platinum; most
preferably platinum.
Embodiment 23
The use according to any one of Embodiment 19-22, wherein the
hydrocracking catalyst comprises at least 0.030 wt %, at least
0.050 wt %, at least 0.075 wt %, at least 0.10 wt %, at least 0.125
wt % or at least 0.20 wt %, of the hydrogenating metal in relation
to the total weight of the catalyst.
Embodiment 24
The use according to any one of Embodiments 19-23, wherein the
hydrocracking catalyst comprises La and/or Ga, preferably at an
amount of 0.10-0.40 wt % of the total weight of the catalyst.
Embodiment 25
The use according to any one of Embodiments 19-24, wherein the
zeolite in the hydrocracking catalyst comprises 70-100 wt % of the
medium pore zeolite and 0-30 wt % of the large pore zeolite with
respect to the total amount of the zeolite.
Embodiment 26
The use according to any one of Embodiments 19-25, wherein the
zeolite in the hydrocracking catalyst comprises 75-95 wt % of the
medium pore zeolite and 5-25 wt % of the large pore zeolite with
respect to the total amount of the zeolite.
Embodiment 27
The use according to any one of Embodiments 19-26, wherein the
hydrocracking catalyst has a deactivation rate of less than
|-3.5.times.10.sup.4 per hour|, preferably less than or equal to
|-3.0.times.10.sup.4 per hour|, or less than or equal to
|-2.5.times.10.sup.4 per hour|, most preferably less than or equal
to |-2.0.times.10.sup.4 per hour|.
Embodiment 28
The use according to any one of Embodiments 19-27, wherein the
medium pore zeolite comprises a ZSM-5.
Embodiment 29
The use according to any one of Embodiments 19-28, wherein the
large pore zeolite comprises a mordenite.
Embodiment 30
The use according to any one of Embodiments 19-29, wherein the
silica to alumina molar ratio of the medium pore zeolite is in the
range of 20-50, preferably 20 to 30, more preferably 21 to 29.
Embodiment 31
The hydrocracking catalyst accordingly to any one of Embodiments
1-30, wherein the hydrocracking catalyst can be free of metals
other than the Group 10 metals of the Periodic Table of Elements,
rhodium, and iridium; preferably free of metals other than
palladium and platinum.
Embodiment 32
The use according to any of Embodiments 19-30 and the process
according to any of Embodiments 1-18, further comprising separating
the benzene from LPG, toluene and xylene, to produce a product
stream, wherein the product stream has a benzene purity of greater
than or equal to 99.80 wt %, preferably greater than or equal to
99.90 wt %, or greater than or equal to 99.95 wt %. It was
unexpected that such a benzene purity could be obtained starting
from a feed stream comprising at least 0.5 wt % of
2,4-dimethylpentane and/or 2,2,3-trimethylbutane (e.g., 0.5 wt % to
15 wt % of 2,4-dimethylpentane and/or 2,2,3-trimethylbutane), based
on a total weight percent of the hydrocracking feed stream.
It is noted that the invention relates to all possible combinations
of features described herein, preferred in particular are those
combinations of features that are present in the claims. It will
therefore be appreciated that all combinations of features relating
to the composition according to the invention; all combinations of
features relating to the process according to the invention and all
combinations of features relating to the composition according to
the invention and features relating to the process according to the
invention are described herein.
It is further noted that the term `comprising` does not exclude the
presence of other elements. However, it is also to be understood
that a description on a product/composition comprising certain
components also discloses a product/composition consisting of these
components. The product/composition consisting of these components
may be advantageous in that it offers a simpler, more economical
process for the preparation of the product/composition. Similarly,
it is also to be understood that a description on a process
comprising certain steps also discloses a process consisting of
these steps. The process consisting of these steps may be
advantageous in that it offers a simpler, more economical
process.
As used herein, deactivation rate is the slope of conversion versus
the time-on-stream (tos). If the conversion is decreasing with time
on stream, the slope is negative. The absolute value of the number
is relevant. The larger the absolute value the faster the catalyst
is deactivating. In other words, a catalyst having a deactivation
rate of |-5.0.times.10.sup.-4 hr.sup.-1| will deactivate 5 times
faster than a catalyst having a deactivation rate of
|-1.0.times.10.sup.-4 hr.sup.-1|. Unless specified otherwise, the
deactivation rate is per hour (hr.sup.-1).
When values are mentioned for a lower limit and an upper limit for
a parameter, ranges made by the combinations of the values of the
lower limit and the values of the upper limit are also understood
to be disclosed.
* * * * *